Note: Descriptions are shown in the official language in which they were submitted.
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PROCESS FOR THE PREPARATION OF LINEAR ALPHA-OLEFINS FROM SYNTHESIS GAS OVER
A COBALT CATALYST
LINEAR ALPHA OLEF1NS FROM NATURAL CiAS-I~ElZ1 V El)
SYNTHESIS GAS OVER A NONSHIFTING COBALT CATALYST
BACKGROUND OF THE DISCLOSURE
FIELD OF THE INVENTION
[0001] The invention relates to producing linear alpha olefins from synthesis
gas over a cobalt catalyst. More particularly the invention relates to
producing
C4-C2o linear alpha olefins having low amounts of oxygenates, by reacting H2
sand CO in a synthesis gas produced from natural gas, over a non-shifting
cobalt
catalyst at reaction conditions of temperature, % CO conversion, Ha:CO mole
ratio and water vapor pressure effective for the mathematical expression of
200 -
0.6T + 0.03PHao - 0.6X~o - 8(HZ:CO) to have a numerical value greater than or
equal to 50. This fan be integrated into a Fischer-Tropsch hydrocarbon
synthesis process producing fuels and lubricant oils.
BACKGROUND OF THE INVENTION
[0002] Linear alpha olefins in the C4-C2o carbon atom'range are large volume
raw materials used in the production of, for example, polymers, detergents,
lubricants and PVC plasticizers. The demand for these olefins is rapidly
increasing, particularly for those that have from 6 to 12 carbon atoms, such
as
six and eight carbon atom linear alpha olefins desirable for making polyolefin
plastics. Most linear alpha olefins are produced by ethene oligomerization,
for
which the ethene feed cost can account for more than half the total cost of
the
alpha olefin production. It is known that alpha olefins can be produced from
synthesis gas using iron, iron-cobalt, iron-cobalt spinet, copper-promoted
cobalt
and cobalt manganese spinet catalysts, most of which are shifting catalysts.
Examples of linear alpha olefin production with such catalysts may be found,
for
example, in U.S. patents 4,544,674; 5,100,856; 5,118,715; 5,248,701; and
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_7_
6,479,557. U.S. 6,479,557, for example, discloses a two-stage process to make
paraffinic hydrocarbons in the first stage.and olefinic hydl-ocarbons LII the
second
stage. The paraffinic .product is made by converting a substoichiolnetric
synthesis gas feed (i.e., H~/CO feed ratio .lower thaIl about 2.1:1) over a
non-
shifting catalyst in the .first stage. Since the I-h/CO usage ratio is
stoichiometric,
the effluent of the first stage is significantly depleted in CO. This effluent
of the
first stage is then used to make olefinic hydrocarL~ons in the second stage
over a
shifting Fischer-Tropsch catalyst.
[0003] Although iron-based shifting catalysts produce hydrocarbons from
synthesis gas with high alpha olefin content even at high CO conversion
levels,
the undesirable water gas shift reaction associated with shifting catalysts
wastes
part (as much as 50°Io) of the CO feed by converting CO to CO2.
Furthermore,
in addition to high CO loss due to the water gas shift conversion of CO to
C02,
iron-based catalysts produce linear alpha olefins containing more than 1 and
even as much as 10 wt% oxygenates. These oxygenates are poisons~to catalysts
used for producing polymers and lubricants from olefins. Hence, the concentra-
tion of oxygenates must be reduced to a level acceptable for polymer and
lubricant production. The methods used for removing the oxygenates are costly,
thus catalysts and processes yielding olefin products with low oxygenate
content
is highly desired.
[0004] It would be an improvement to the art if a way could be found to (i)
produce linear alpha olefins with low oxygenates levels and particularly with
(ii)
a non-shifting catalyst and preferably a non-shifting Fischer-Tropsch hydro-
carbon synthesis catalyst that is also useful for synthesizing fuel and
lubricant oil
fractions. It would be a still further improvement if (a) linear alpha olefin
production could be integrated into a Fischer-Tropsch hydrocarbon synthesis
process and (b) if a hydrocarbon synthesis reactor employing a non-shifting
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Fischer-Tropsch hydrocarbon synthesis catalyst and producing fuel and
lubricant
oil fraction hydrocarbons could also be used for linear alpha olefin
production,
and vice-versa, without having to change the catalyst in the reactor.
SUMMARY OF THE INVENTION
[0005] The invention relates to a process for producing linear alpha olefins,
and particularly linear alpha olefins having from four to twenty carbon atoms,
having less than 3 and preferably less than 1 wt% oxygenates, by reacting Ha
and CO, in the presence of a non-shifting Fischer-Tropsch hydrocarbon
synthesis catalyst comprising a catalytic cobalt component, under reaction
conditions defined by a Condition Factor (CF) greater than or equal to 50,
which
Condition Factor is defined as:
i
CF = 200 - 0.6T + 0.03PHZO - 0:6Xco - ~(HZ:CO)
where,
T = average reactor temperature in °C; the average reactor
temperature is
calculated by averaging the temperature readings from thermocouples
measuring the temperatures of individual segments of the reactor.
For example, if the temperature is measured in the middle of the first,
second, and third equal-volume segments of a fixed bed reactor, the
average temperature is equal to one third of the sum of the three
readings.
PH20 = p~'tial pressure of the water in the synthesis gas feed to the
reactor, in kPa; the partial pressure of water in the feed is
calculated by multiplying the mol friction of water in the feed by
the feed pressure measured in kPa. The mol fraction of feed
components can be determined by, for example, using gas
chromatographic methods.
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X~o = CO conversion expressed as percent; the CO conversion can be
determined from the CO balance. There are many methods
available for establishing material balance. The method herein
utilized measurements based on an internal standard such as a
noble gas or nitrogen that is inert during Fischer-Tropsch
synthesis. When using an inert internal standard, the conversion
can be simply calculated by measuring the concentrations of CO
and the internal standard in the feed and the effluent. This and
other calculation methods are well known in the art of chemical
engineering. The concentrations of CO and the inert internal
standard in turn can be determined by gas chromatographic
methods known in the art.
1'12:C0 = HZ to CO molar ratio in the synthesis gas feed to the reactor; the
concentrations of H2 and CO in the feed can be determined by
gas chromatographic methods. . ,
[0006] By nonshifting is meant that under the reaction conditions the catalyst
will convert less than 5 and preferably less than 1 mole % of the CO to COZ up
to
90% CO conversion in Fischer-Tropsch synthesis. The wt% of oxygenates is
meant the wt% of oxygenates in the total synthesized C4-C2o hydrocarbon
fraction, and by oxygenates is meant oxygen-containing hydrocarbon molecules,
such as alcohols, aldehydes, acids, esters, ketones, and ethers. The process
of
the invention has been found to produce a C4-C2o hydrocarbon fraction contain-
ing more than 50 wt% linear alpha olefins and less than 3 wt% preferably less
than 1 wt% oxygenates. This process can be achieved as a stand-alone process
or it can be added to or integrated into a Fischer-Tropsch hydrocarbon
synthesis
process. The relatively low selectivity for alpha olefin production normally
exhibited by a non-shifting Fischer-Tropsch cobalt catalyst, is at least
partially
overcome by operating the hydrocarbon synthesis reactor under reaction
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conditions in which the CF, according to the above expression, is greater than
or
equal to 50.' By CO conversion is meant the amount of CO in the synthesis gas
feed converted in a single pass through the reactor.
[0007] In another embodiment the invention relates to (a) producing a CO
and HZ containing synthesis gas from natural gas, (b) reacting the H2 and CO
containing synthesis gas in the presence of a non-shifting cobalt Fischer-
Tropsch
hydrocarbon synthesis catalyst, at reaction conditions effective to achieve a
Condition Factor (CF) greater than or equal to 50, to synthesize linear alpha
olefins, and particularly linear alpha olefins having from four to twenty
carbon
atoms, having less than 3 wt%, preferably less than 1 wt% oxygenates. A
process in which natural gas is converted to synthesis gas which, in turn, is
converted to hydrocarbons, is referred to as a gas conversion process. In yet
another embodiment, the process of the invention relates to an integrated .gas
conversion process, in which the linear alpha olefin production process of the
invention is integrated with a hydrocarbon synthesis process which produces
primarily fuel and lubricating oil products. This is explained in detail
below.
[0008] It is preferred in the practice of the invention that the synthesis gas
be
produced from natural gas. Natural gas typically comprises mostly methane, for
which the H:C .ratio is 4:1 and is therefore an ideal feed for producing
synthesis
gas having a nominal H2:C0 mole ratio of 2:1 or somewhat higher, for example
2.1:1. Substantial amounts of hydrogen can be separated from a synthesis gas
with a HZ:CO = 2:1 mole ratio, to produce H2 and a 1:1 H2:C0 mole ratio
synthesis gas. The 1:1 H2:C0 mole ratio is a preferred ratio for the linear
alpha
olefin process of the invention. Thus, while the synthesis gas produced in the
synthesis gas generating reactor of a gas conversion plant typically has an
H2:C0
mole ratio of 2.1:1 or 2:1, all or a portion of this synthesis gas may be
optionally
treated to change the H2:C0 mole ratio in the gas to a more preferred ratio
for
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the alpha olefin synthesis process, before it.is passed into the one or more
linear
alpha olefin producing reactors.
[0009] It is understood that although the chemistry involved in producing
linear alpha olefins over a non-shifting Fischer-Tropsch cobalt catalyst is
such
that it is preferred for the HZ:CO mole ratio in the synthesis gas feed passed
into
the linear alpha olefin reactor be typically less than 2:1, the stoichiometric
H2:C0 consumption mole ratio of the linear alpha olefin reaction is 2:1.
Furthermore, it is also understood that conventional hydrocarbon synthesis
making paraffinic hydrocarbons for fuel and lubricant applications over a non-
shifting cobalt catalyst employs a synthesis gas feed in which the HZ:CO mole
ratio is 2.1:1. In an integrated gas conversion process of the invention, one
or
more rectors may be added to and/or switched back and forth from hydrocarbon'
synthesis for producing fuel and lubricant fractions to linear alpha olefin
production by changing the conditions from the conventional hydrocarbon
synthesis conditions to the alpha olefin selective conditions of the present
invention. by adjusting the reaction parameters to achieve a CF value of
greater
than or equal to 50 and vice versa. Thus, producing the synthesis gas from
natural gas provides a particular synergy and flexibility for practicing all
embodiments of the present invention.
[0010] A CF value of greater than or equal to 50 can be achieved by many
different combinations of temperature, CO conversion, H2:C0 ratio and water
partial pressure. The individual process conditions that are preferred for
achieving CF greater than or equal to 50, and thus high alpha olefin
selectivity
and concomitant production, include (a) setting the H2:C0 ratio in the
synthesis
gas feed to a value less than 2.1:1 and preferably less than 1.8:1, (b) a CO
feed
conversion of less than 50 and preferably less than 30% in a single pass
through
the reactor, and (c) a reaction temperature typically between 160 and 250 and
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_7_
preferably between 180 and 240°C. The presence of water in the
synthesis gas
feed, while optional, is preferred. Thus, it is possible for the value for
PH~o in
the expression above to be zero or negligible but a non-zero value in the
range of
from 50 to 500 kPa is preferred.
[0011] It is understood that although the above-specified preferred ranges
provide guidance for the typical values of the individual control variables,
(i.e.,
average reactor temperature, partial pressure of water and HZ:CO ratio in the
gas
feed to the reactor, and CO conversion) under which a Condition Factor value
of
greater than-or equal to 50 can be achieved, it is the combination of the
control
variables which satisfy the CF described herein which is the invention not the
individual variables alone. Thus the skilled artisan should appreciate that
the
invention is a linear combination of the control variables as defined by the
Condition Factor. 'The utility of the Condition Factor is that it enables the
determination of this preferred combination of the control variables. Thus for
example, if for economic or process reasons the feed to the reactor does not
have
steam (i.e., Pro = 0 kPa), and the CO conversion needs to be at least 30%, the
H2:C0 ratio in the feed to the linear alpha olefin reactor and the average
reactor
temperature need to be set so that the sum of 0.6(T) and 8(HZ:CO) is less than
200-18-50 = 132. If, again, for process reasons the temperature is set for
205°C, ,
the H2:C0 ratio needs to be less than 1.125:1 for the process of the
invention'.
Clearly, if three control variables are set for process or economic reasons,
the
preferred value range of the fourth variable can be readily calculated. Those
skilled in the art will also recognize, that if only two control variables are
fixed,
the preferred combinations of the remaining two variables will define a two-
dimensional surface which will be further defined by some other common
boundary conditions known in the art, like that the H2:C0 ratio cannot be
equal
or less than zero (no hydrocarbons form in the absence of H2) or that the
reaction
temperature cannot be below 160°C (where for the catalyst systems of
interest to
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rocess of this invention, the Fischer Trapsch reaction rate is too low.ta be
'tee P
practical). Likewise, if only one control variable is f xed, the preferred
condi-
tions define a three-dimensional space. It is c~.ear, therefore, that yvh~le
the
specific pref~errred ranges given earlier for the individual control variables
provide reasonable starting points, the ultimate combination of the control
variables of the invention need to be derived from the expression provided
herein 'far 'tee Condition Factor.
~OOlz] 'Thus, although Applicants have specified preferred ranges for the
input variables for CF (T, Pro, ~co ~ Hz:CO) sa long as the specified CF is
met; the input variables may vary from the ranges specked herein. Hence the
CF criteria disclosed herein, is the critena'~''~hich must be met when setting
the
noted input variables.
(001.3] Lower C~ conversion can be readily achieYed by increased synthesis
gas feed rates through the xeactor, which also results in shorter product
residence
times. As a consequence, synthesis gas feed rate is another variable 'that may
be
used tv achieve the desired conversion level and thus a CF &Teate~r than. or
equal
tv 50. Thus, the synthesis gas feed rate (commonly quantified as Gas Hourly
Space Velocity or GHfV) through the reactor, with a n~on-shi~g Catalyst such
as disclosed inUS °5,945,459, US 5,968,991, US 5,090,742, U5 6,136,5689
LJS
6,319,960,.US RE 37,406E, US 6,355,593, US 6,331,575 campnsing a catalytic
cobalt component; will typically be greater than 15,000 standard volumes of
gas
(measured at 103 kFa and ~5°G)/volume of catalystlhour (V~~T), ~d
preferably geeater than 25,000 VfV/hr. It is understood, however, that the
feed
rate necessary. to maintain ~'e p~'efe~ed C4 conversion of the invention will
also
a d on the volumetric productivity of the catalys H ce, as wou d be~~ .
dep n
readily appas.'ent to the skilled artisan, a catalyst that has two tunes
higher
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CA 02507722 2005-05-30
WO 2004/060836 PCT/US2003/036896
volumetric activity, will require two times faster feed rate to maintain the
same
CO conversion or 50,000 vlv/hr.
[0014] These conditions, particularly the low CO feed conversion and high
synthesis feed gas rate through, the reactor, are more readily achieved in a
reactor
containing one or more fixed beds of catalyst or a fluidized catalyst. A
slurry
reactor, highly efficient for synthesizing higher molecular weight paraffinic
hydrocarbons, may also be utilized for the olefin synthesis, provided that the
appropriate residence times are maintained. Such residence times are readily
'determined byrtlle skilled artisan due to the inherently longer product
residence
time.
[0015] In order to maximize the hydrocarbon product yield in the C4-C2o
carbon range, particularly the linear alpha olefin yield in the C4-C2o carbon
range, the hydrocarbon synthesis reaction is preferably conducted at an alpha
of
less than 0.9 and more preferably less than 0.8. This is in contrast to an
alpha of
at least, and preferably greater than 0.9, which is desirable for synthesizing
higher molecular weight hydrocarbons for fuel and lubricant applications.
[0016] In a broad embodiment, the invention relates to a process for
synthesizing CQ-CZO linear alpha olefins, wherein the process comprises
passing a
synthesis gas feed comprising a mixture of HZ and CO through a linear alpha
olefin hydrocarbon synthesis reactor, in which it contacts a non-shifting
Fischer-
Tropsch hydrocarbon synthesis catalyst comprising a catalytic cobalt
component, at reaction conditions sufficient for the H2 and CO to react and
form
the linear alpha olefins and wherein the reaction conditions are such that the
following expression has a value greater than or equal to 50:
200 - 0.6(T) + 0.03PHao - 0.6Xco - 8(H2:C0),
and wherein,
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T = average reactor temperature in °C
PH20 = p~'tial pressure of the water in the synthesis gas feed to the
reactor, in kPa
X~o = CO conversion, 'expressed as percent, and
H2:C0 = hydrogen to CO molar ratio in the synthesis gas feed to the
reactor.
[0017] The above expression defines the Condition Factor (CF). Thus, the
preferred-reaction conditions of the process of the invention can also be
described as a combination of average reaction temperature, water partial
pressure in the feed, CO conversion, and feed H2:C0 ratio that yields a CF
value
of greater than or equal to 50.
BRIEF DESCRIPTION OF THE DRAWINGS
[0018] Figure 1 is a block flow diagram of a stand-alone linear alpna.olefin
process of the invention.
[0019] Figure 2 is a block flow diagram of one embodiment of an integrated
linear alpha olefin and hydrocarbon synthesis process of the invention.
[0020] ~ Figure 3 is a block flow diagram of one embodiment of another
integrated linear alpha olefin and hydrocarbon synthesis process of the
invention.
DETAILED DESCRIPTION
[0021] In the process of the invention, the linear alpha olefins are produced
by a Fischer-Tropsch hydrocarbon synthesis reaction in which H2 and CO in the
feed gas react in the presence of a non-shifting Fischer-Tropsch hydrocarbon
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synthesis catalyst, comprising a catalytic cobalt component, under reaction
conditions defined by a Condition Factor described above as having a value
greater than or equal to 50. Optionally, the process conditions are also
adjusted
and/or the catalyst selected to achieve a reaction alpha of less than 0,9 and
preferably less than 0.8 to maximize linear alpha olefin production 'in the C4-
CZo
carbon range. By reaction alpha is meant the Schultz-Flory alpha, as is deter-
mined by the molecular weight distribution of the synthesized hydrocarbons,
and
may be determined as reported in J. Eilers, et. al., The Shell Middle
Distillate
Synthesis Process (SMDS) in Catalysis Letters Vol. 7 (1990) p. 253-270. At a
given H2 to CO mole ratio; the alpha will be reduced by (a) increasing the
reaction temperature and (b) decreasing the reaction pressure.
[0022] The non-shifting hydrocarbon synthesis catalyst comprising a catalytic
cobalt component Bused in the linear alpha olefin production process of the
invention comprises unsupported cobalt, or a supported Fischer-Tropsch
hydrocarbon synthesis catalyst which is a composite of cobalt or promoted
cobalt and one or more support components. The preparation of an unsupported
or bulk, rhenium-promoted cobalt catalyst that has been found useful for
producing linear alpha olefins according to the invention is described below,
in
the preamble to the examples. Also described below, and which is useful for
producing linear alpha olefins according to the invention, is a supported,
rhenium-promoted cobalt catalyst, that has been used to produce linear alpha
olefins according to the process of the invention and has also been used in
slurry
Fischer-Tropsch hydrocarbon synthesis reactors producing mostly (e.g, at least
90 wt%) saturated, normal paraffin hydrocarbons having less than 6 wt% olefins
(less than 10 wt% olefins in the synthesized C4-C2o fraction), wherein the
support comprised a composite of titanic and silica-alumina. Preferred
supports
for supported Fischer-Tropsch hydrocarbon synthesis catalysts containing a
catalytic cobalt component, comprise titanic, silica, and modified titanic and
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silic such as, for exa~nnple, e.~4z-modified TiOz, with or without binders,
for
both the linear alpha olefin synthesis anal for synthesising higher molecular
vv~eight, primarily para~c li9.m'd hydrocarbon products, including lubricant
oil
fractions. A support component comprising tita~a ~ preferred. The amvun~ of
cobalt present in suppvrked catalysts may mange from 1-50 wt°Aa of the
catalyst,
preferably 2-40 wt°/a and more preferably 2-25 wt°/4. If a
promoter such as V
rhenium is used, the weight ratio of the cobalt to the promoter will range
from.
30:1 to 2:1 and preferably froia 20:1 to 5:1. Useful catalysts and their
prepara-
tion are lmown and illustrative, but non-limiting examples may be found, for
example, iu US $,945,459, US 5,968,991, US 6,090,742, US 67136,865, US
6,319,960, US RE 37,406E, US 6,355,593, US 6,331,575.
jpp231 The total synthesized C4-Czo hydrocarbon fraction is separated by
fractionation, from the lower and higher carbon number, C3_ and Gzm
hydr°'
~bons, which are also synthesised by the linear alpha olefin producing process
of the inventive. The separated C4-Coo hydrocarbon fraction containing the
desired linear alpha olefins also contains internal olefins, paraffins, and a
small
amount of oxygenates formed by the synthesis reaction. It is therefore typical
to
further treat the linear alpha olefins to recover the desired C4-C2a bear
alpha
olefins, by any suitable means. All known linear alpha olefin recovery
processes
are quite complex and the caStlY- It is therefore advantageous to make a crude
alpha olefin product that has as high a concentration of the linear alpha
olefins as
possible, to thereby reduce the complexity anal cast of recovering them- Ome
known method for recovering linear alpha olefins from the oxygenates-
. . containing, total, Synthesized C~-C,~o hydrocarbon fracti.vn, is to react
the linear
alpha olefins in the separated G~-Czo fxaction with alkanols, to form ethers.
The
so-formed ethers are then 9ep~ted from the rest of the hydrocatbDns in the
separated C4-Coo Gut by fractionation (c.f. German patent publications DE
19825295 and DE 19833941). 'The separated ethers are chemically treated to
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CA 02507722 2005-05-30
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-13-
convert them back into linear alpha olefins. Since ethers are themselves
oxygenates, the presence of oxygenates in the separated total C4-C2o
hydrocarbon fraction increases the difficulty and cost of recovering the
linear
alpha olefins by this process. Therefore, a significant advantage of the
process
of the present invention is the very low levels of oxygenates of less than 3
wt%
and preferably less than 1 wt% that it produces in the total C4-Cao
hydrocarbon
fraction synthesized by the process. This low oxygenates level is much lower
than the typical amount of at least 8-10 wt% oxygenates made using iron-based
Fischer-Tropsch hydrocarbon synthesis catalysts. The total oxygenate content
may be determined by gas chromatography or by high-resolution NMR.
Another known method for separating linear alpha olefins from saturated
hydrocarbons and internal olefins in the same boiling point range may be
found,
for example, in U.S. patent 5,877,378. In this process, which is also
adversely
i
affected by the presence of oxygenates (e.g., alcohols and acids); the linear
alpha
olefins are selectively converted to trialkylaluminum compounds, which are
then
separated from the unconverted other hydrocarbons and converted back into
ethers. The low oxygenates content and high level of linear alpha olefin
production obtained in the practice of the invention, results in significantly
lower
cost to recover the linear alpha olefin products by this method as well.
[00241 The linear alpha olefin producing process of the invention can be
integrated into a Fischer-Tropsch hydrocarbon synthesis process plant
(integrated process) or it may comprise a separate facility with its own
source of
synthesis gas (stand-alone process). As mentioned under the summary, it is
preferred that all process embodiments comprise gas conversion processes,
which include production of the synthesis gas feed from natural gas. Syngas
derived from natural gas provides the desired stoichiometry for the linear
alpha
olefin synthesis process of the invention which utilizes a non-shifting cobalt-
based Fischer-Tropsch synthesis catalyst. The linear alpha olefin production
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-14-
process of the invention is also a type of Fischer-Tropsch hydrocarbon
synthesis
process, but one that maximizes olefin production to at least 50 wt% of the
total
synthesized hydrocarbons, i.e., at least 50 wt% of the total synthesized C4-
C2o
hydrocarbons produced in the linear alpha olefin reactor are linear alpha
olefins.
This distinguishes it from what is referred to herein as a hydrocarbon
synthesis
process, which produces mostly (e.g., at least ~5 wt%) saturated, i.e.,
paraffinic
hydrocarbons overall, with typically less than 5 vt% linear alpha olefins in
the
synthesized Cø-C2o hydrocarbon fraction. Therefore, for the sake of
convenience
and in the context of th'~ invention, the latter process is referred to herein
as a
hydrocarbon synthesis process, while the process described herein as having a
CF greater than or equal to 50 is referred to as a linear alpha olefin
synthesis
process. ,
[0025] In an integrated linear alpha olefin and hydrocarbon synthesis (HCS)
process of the invention, synthesis gas is fed to either the LAO or HCS
reactor.
The tail gas from~th~e reactor can then be utilized as feed gas for the other
reactor. Additionally, products from one of the two processes (hydrocarbon
synthesis and linear alpha olefin synthesis) may be combined with products
from, the other process.
[0026] Preferably in an integrated process syngas having e.g. an H2:C0 mole
ratio of 2:1 is passed to an HCS reactor and syngas having an H2:C0 mole ratio
of lower than 2:1 is passed to an LAO reactor because the consumption ratio of
the LAO process is 2:1 the tail gas from the LAO reactor will be CO rich and
H2
poor or depleted. The tail gas from both units can then be combined to provide
a
feed gas to the LAO reactor having a more preferred HZ:CO mole ratio of lower
than 2:1. In this manner, an integrated process is achieved. Alternatively,
the
tail gas from the HCS reactor can be used alone as feed gas to the LAO
reactor.
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[0027] One or more reactors for producing linear alpha olefins according to
the practice of the invention may be erected as part of a hydrocarbon
synthesis
plant, added to it later or be one or more hydrocarbon synthesis reactors that
have been temporarily or permanently switched to linear alpha olefin
production.
In another integrated linear alpha olefin process embodiment, in a plant or
facility (these two terms are used synonymously) built for linear alpha olefin
production, one or more hydrocarbon synthesis reactors may be added to it or
one or more linear alpha olefin producing reactors may be temporarily or
permanently switched to hydrocarbon synthesis reactor. In an embodiment in
which a hydrocarbon synthesis reactor contains anon-shifting synthesis
catalyst
comprising a catalytic cobalt component, replacement of the catalyst will not
be
necessary when switching it over from hydrocarbon synthesis to linear alpha
olefin synthesis providing a.unique benefit. Further, when switching a linear
alpha olefin synthesis reactor to hydrocarbon synthesis, it will not be
necessary
to change the catalyst in the reactor for the switchover.
[0028] In one specific embodiment of an integrated linear alpha olefin
synthesis process of the invention, one or more linear alpha olefin synthesis
reactors will use tail gas from one or more Fischer-Tropsch hydrocarbon
synthesis reactors, as all or part of the synthesis gas used for producing the
linear ,
alpha olefins and vice-versa. By way of an illustrative, but nonlimiting
example
of this type of embodiment, at least two synthesis reactors are used, with at
least
one reactor, a Fischer-Tropsch hydrocarbon synthesis reactor upstream of the
linear alpha olefin reactor, operating with a non-shifting cobalt hydrocarbon
synthesis catalyst comprising a catalytic cobalt component and a support
component, at a reaction alpha high enough (e.g., >_ 0.9) to produce both fuel
and
lubricant oil hydrocarbon fractions, from a synthesis gas feed (preferably
produced from natural gas) in which the HZ:CO mole ratio is less than the
stoichiometric (2.1:1) H2:C0 consumption ratio (e.g., 2.0:1 or less) and the
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reactor is operating at a high CO conversion level (e.g., 80% or higher). The
hydrocarbon synthesis reactor may be a slurry reactor, a fixed bed reactor or
a
fluid bed reactor. A slurry reactor is preferred for maximizing higher
molecular
weight hydrocarbons boiling in the lubricant oil range. Both the linear alpha
olefin and hydrocarbon synthesis reactors of the integrated linear alpha
olefin
process may contain the same or different non-shifting hydrocarbon synthesis
catalyst comprising a catalytic cobalt component: The hydrocarbon synthesis
reactor produces a reactor tail gas containing unreacted H2 and CO, having a
reduced H2:C0 mole ratio. After removing at least a portion of the Ca+ hydro-
carbons and (optionally) H20 from the tail gas, the remainder containing the
reduced (e.g., l:l) mole ratio H2:C0, is passed into the linear alpha olefin
synthesis reactor. The conditions of the linear alpha olefin synthesis reactor
is ,
set so that the CF value, which incorporates the aforementioned reduced feed
ratio, is higher than 50. In yet another illustration of an integrated process
embodiment of the invention, a synthesis gas producing unit is producing, from
natural gas, a synthesis gas containing H2 and CO in a mole ratio of at least
2:1,
as feed for one or more hydrocarbon synthesis reactors. A portion or slip
stream
of this synthesis gas is passed to either physical or chemical separation
means,
for removing some of the hydrogen from the synthesis gas, to produce a
hydrogen-reduced synthesis gas in which the H2:C0 mole ratio is less than
2.1:1
(e.g., l:l). This hydrogen-reduced synthesis gas is passed into one or more
linear alpha olefin synthesis reactors. The conditions of the linear alpha
olefin
synthesis reactor is again set so that the CF value is greater than or equal
to 50.
Physical separation processes (means) for separating H2 from the synthesis gas
include adsorption-desorption, membrane separation and a combination thereof,
all of which are well known and commercially available. Adsorption-desorption
processes include temperature swing adsorption (TSA) and pressure swing
adsorption (PSA), both of which comprise a plurality of adsorbent containing
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vessels operated in a cyclic manner. Chemical means includes a water gas shift
reactor, which is typically combined with physical separation means.
[0029] Synthesis gas for the process of the invention is preferably produced
from natural gas, which can comprise as much as 92 mole % methane, with the
remainder primarily Cz+ hydrocarbons, nitrogen and C02. Methane has a 4:1
H:C ratio and is therefore ideal for producing, preferably by a combination of
partial oxidation and steam reforming, a synthesis gas having an H2:C0 mole
ratio of nominally 2.1:1 a which is the stoichiometric mole ratio used with a
non-
~shifting cobalt catalyst for hydrocarbon synthesis. Sulfur and other
heteroatom
compounds are removed from the natural gas, and in some cases also nitrogen
and C02. The remaining methane-rich gas, along with oxygen or air and steam,
is passed into a synthesis gas generator. Oxygen is preferred to air, because
it
does not introduces nitrogen into'the synthesis gas generator (reactor).
During
the synthesis gas reaction, nitrogen can form HCN and NH3, both of which are
poisons to a cobalt Fischer-Tropsch catalyst and must therefore be removed
down to levels below 1 ppm. If nitrogen is not removed from the natural gas,
and/or if air is used as the source of oxygen, HCN and NH3 must be removed
from the synthesis gas, before it is passed into a hydrocarbon synthesis
reactor.
Known processes for synthesis gas production include autothermal reforming
and fluid bed synthesis gas generation, both of which employ oxygen and form
the synthesis gas by partial oxidation and catalytic steam reforming. A review
of
these and other processes for producing synthesis gas and their relative
merits,
may be found, for example, in U.S. patent 5,883,138.
[0030] Due to the nature of the surface chemistry involved, hydrogen to
carbon monoxide mole ratios of typically less than 2:1 are preferred
preferably
less than 1.8:1, more preferably about 1:1 in the feed to the reactor for
achieving
a preferred CF value of greater than or equal to 50 and thus producing linear
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alpha olefins with high selectivity over a non-shifting synthesis catalyst
compris-
ing a catalytic cobalt component. However, the stoichiometric H2:C0 mole ratio
of the H2 and CO consumed by the linear alpha olefin synthesis reaction is 2:1
and the consumption of the CO in the synthesis gas feed gas by the reaction in
a
single pass through the synthesis reactor is typically less than 50%. This
means
the linear alpha olefin synthesis gas reactor produces a tail gas which (i) is
rich
in valuable, unreacted CO, (ii) depleted in HZ and (iii) has an H2:CO mole
ratio
below that in the feed gas passed into the reactor. At least a portion of this
tail
gas, comprising the unreacted CO and hydrogen, is recycled back into the
linear
alpha olefin synthesis reactor, along with the incoming, fresh synthesis gas
feed.
The fresh makeup gas fed to the process will typically have a higher H2:C0
ratio
than the H2:C0 ratio in the feed to the reactor if such recycle of the linear
alpha
olefin synthesis reactor effluent gas is applied, and will essentially follow
the
stoichiometry of the linear alpha olefin synthesis process, i.e., typically
2:1. It is
important to understand that in cases when there is an effluent recycle, the
HZ:CO ratio used in calculating the Condition Factor is not that of the make-
up
gas but instead the H2:C0 ratio of the feed to the linear alpha olefin
synthesis
reactor. The feed to the reactor is generated by blending the recycle and make-
up gas streams, thus the reactor feed typically has lower than 2:1 H2:C0 ratio
due to the H2-deficiency of the recycled tailgas component. Hydrogen may be
separated from synthesis gas having a relatively high H2:C0 mole ratio of,
e.g.,
2:1, or 2.1:1, to produce hydrogen and a synthesis gas having a lower Ha:CO
mole ratio. One or more of the higher or lower mole ratio synthesis gases, and
separated hydrogen, may be combined with the recycled CO-rich tail gas, to
provide the desired H2:C0 mole ratio in the combined synthesis gas stream fed
into the linear alpha olefin synthesis reactor. It is understood, however,
that
regardless of the feed HZ:CO ratio to the reactor, the Ha:CO comsumption ratio
of the linear alpha olefin process of the invention is 2:1. This consumption
ratio
is close to the 2.1:1 ratio needed for the conventional hydrocarbon synthesis
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process carried out with non-shifting cobalt-based catalysts making paraffinic
fuel and lubrication products. As discussed earlier, the H2:C0 ratio in
synthesis
gas derived from natural gas is also 2:1, and thus in line with the
consumption
ratio of both the conventional hydrocarbon synthesis and the linear alpha
olefin
synthesis processes. Consequently, the use of natural gas to produce synthesis
gas is preferred for both linear alpha olefin process of the invention and
those
embodiments in which it is integrated with hydrocarbon synthesis.
[0031] Fischer-Tropsch hydrocarbon synthesis processes are well known and
~-~-omprisewcontacting'a synthesis gas comprising ~a mixture of H2 and CO with
a
Fischer-Tropsch synthesis catalyst, at reaction conditions effective for the
HZ and
CO react to form hydrocarbons under shifting or non-shifting conditions. For
the linear alpha olefin synthesis process of the invention, non-shifting
conditions
and a non-shifting catalyst are a~ed. This means that less than 5 and
preferably
less than 1 wt% of the CO in the synthesis gas feed is converted to COZ in one
pass of the synthesis gas through the reactor. Fischer-Tropsch types of
catalysts
for hydrocarbon synthesis may comprise, for example, one or more Group VIII
catalytic metals such as Fe, Ru, Co, and Ni, and optionally one or more
promoters such as Re, Pt, Th, Zr, Hf, LT, Mg and La, on a suitable inorganic
support material, preferably one that comprises one or more refractory metal
oxides. A non-shifting catalyst in which the catalytic metal component
comprises cobalt is used in the linear alpha olefin reactor of the invention
and
this typically excludes the use of iron in the catalyst. Some of the
synthesized
hydrocarbons will be liquid, some solid (e.g., wax) and some gas at ambient
conditions of temperature and pressure, particularly if a catalyst having a
catalytic cobalt component is used. Slurry Fischer-Tropsch hydrocarbon
synthesis processes are often preferred for producing relatively high
molecular
weight, paraffinic hydrocarbons when using a cobalt catalyst, but not for
producing the linear alpha olefins. The H2:C0 mole ratio for a hydrocarbon
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synthesis process may broadly range from about 0.5 to 4, but is more typically
falls within the range of from about 0.7 to 2.75. These H2:C0 mole ratio
ranges
are~achieved by the type of reaction and reaction conditions used to produce
the
synthesis gas.
[0032] Figure 1 is a simple block flow diagram of a stand-alone linear alpha
olefin process of the invention. Plant 10 comprises a fixed bed, down-flow
linear alpha olefin reactor 12, which synthesizes C4-C2o linear alpha olefins
(and
also other hydrocarbons such as C4-C2o paraffins and internal olefins, C3_ and
Cai+ hydrocarbons, etc.) and separation units 14 and 16. Synthesis gas compris-
ing H2 and CO is passed, via line 18, from a synthesis gas generating unit
(not
shown) which includes gas clean-up to remove sulfur, HCN and NH; from the
gas, into~reactor 12 which contains one or more fixed beds of the non-shifting
Fischer-Tropsch catalyst. In this illustration, the catalyst comprises
unsupported,
rhenium promoted, particulate cobalt. In reactor 12, the H2 and CO in the
synthesis gas react in the presence of the catalyst to form linear alpha
olefins,
including C4-C2o linear alpha olefins (and also saturated paraffins). Reactor
12
operates at reaction conditions that meet the criteria for the Condition
Factor
(CF) to be greater than or equal to 50 which, in this non-limiting
illustration,
includes a synthesis gas H2:C0 mole ratio of 1:1, a CO conversion of 12-15%, a
temperature of 205°C, a synthesis gas feed space velocity of 38,000
V/V/hr, a
pressure of 2000 kPa. The synthesized hydrocarbons and unreacted synthesis
gas pass down out of the reactor and are removed at the bottom by line 20. The
mixture of hydrocarbons and gas passes into separation unit 14, in which the
gas
is separated from the hydrocarbons and the hydrocarbons are separated into a
CZO_ fraction and a CZO+ fraction. The unreacted synthesis gas is removed from
14 via line 26 and recycled back into hydrocarbon synthesis reactor 12, via
line
18, in which it mixes with the fresh synthesis gas feed coming from the
synthesis
gas generating unit. Some of the gas is purged via line 28, to prevent build-
up of
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normally gaseous hydrocarbons (e.g., C~-Cø) in the recycle gas. The Hz:CO
mole ratio~in the synthesis gas produced in the synthesis gas generator is
adjusted so that the mixture of the fresh synthesis gas and recycle gas,
entering
reactor 12, contains HZ and CO in the desired H2:C0 mole ratio which, in this
illustration, is 1:1. Separation unit 14 contains the necessary coolers and
separator drums to condense the water and hydrocarbons produced by the
synthesis reaction and to and separate the water, unreacted synthesis gas and
gaseous hydrocarbons, and liquid hydrocarbons. If desired or necessary, the
separated water is used to humidify the feed gas in line I8, to increase the
partial
pressure of the water in the feed being passed into reactor 12. The liquid
hydro-
carbons are also separated in 14 by fractionation into the C21+ hydrocarbons
and
the linear alpha olefin-containing C2o_ fraction. The CZO_ fraction is removed
from 14 via line 22 and passed into linear alpha olefin separator 16. ~ The
CZo-,.
i
fraction is removed from 14 via line 24. In unit' 16, the linear alpha olefins
are
separated from the other components of the C2o_ hydrocarbon fraction. There
are
several methods known in the art for recovering linear alpha olefins from
Fischer-Tropsch synthesized hydrocarbon streams and any of these methods is
applicable for the recovery of the product C4-C2o alpha olefins of the
invention.
Typically, the linear alpha olefins are selectively converted into another
entity,
like ethers (c.f., German patent publications DE 19825295 and DE 19833941) or
metal alkyls, like alkyl aluminum (US 5,877,378). These chemically
transformed linear alpha olefin derivatives in turn are separated from the
rest of
the unconverted CZO_ hydrocarbons, such as paraffins and internal olefins,
etc., by
fractionation. Finally, the separated linear alpha olefin derivatives are
chemically converted back into linear alpha olefins. The linear alpha olefin
products are removed from the linear alpha olefin recovery unit 16 via line 30
and the rest of the C2o_ hydrocarbons via line 32, from which they are passed
into
line 24, in which they are combined with the C2i+ hydrocarbons. Line 24 passes
these combined C21+ and CZO_ hydrocarbons to further processing.
~ ii.. ( k ~ t
i,xp ~ i, », CA 02507722 2005-05-30 's
f
. ~ # .. i.s 1 f. i L ~ a ~ f n,.M H:
~ ~ E ~ ~~:~, _ ~ Uao33~~~~ ~,~
~tov-09-2004 04:Z5pm From-EXXONMOBIL LAW DEPT . 909-T30-3fi49 T-711 P.OOBl014
F-600
-2Z-
[0033] In Figure Z, an integrated linear alpha olefin synthesis unit 50
Comprises a slurry hydrocarbon synthesis reactor 52, a linear alpha olefin.
Synthesis reactor 12, and separation units 14 and 16. Reactor 12 and units 14
and 16 are the same as in Figure 1, as are the flaw lines having 'the same
numbers as in. Figure 1. Slurry hydrocarbon synthesis reactor 52 contains a
hydrocarbon slurry within in which. is dispersed a particulate Fischer-Tropsch
hy~oc~on catalyst campzising rhenium-promoted cobalt on a titanic support
~e-g. 11 wt% Co and 1 wt°fo Re, based an the total. catalyst weight).
Examples of
"seful catalysts fox hydrocarbon synthesis axe described, for example, in. US.
t
945,459, US 5,968,991, US 6,090,742, US 6,136,865, US 6,319,960, US R'~
37,406E, US 6,355,$93, tJS 6,331,5'15. In this embodiment, a spathesls gas
feed
compri$ing H~ and CO in a Z.1:1 mole ratio is passed up mto reactor 52 ~.a ~.e
54. Reactor 5Z operates at an alpha of >_ 0.9 and produces mostly te-g-= ~ ~0
.~~'o para~.ns), including hydrocarbons boiling in the fuels and lubricant
ranges.
The CO conversion in the reactor is 80 °t° in a single pass. The
synthesized
hydrocarbons that are liquid at the reaction conditions are separated from the
catalyst particles in 52 via. filtration (not shown) and removed from the
reactor
via line Sfi. The utsreacted synthesis gas, which now contains Hz and CO in a
lower than 2: I ,mole ratio (preferred for linear alpha olefinsynthesis), is
removed
$om the reactor via line 5$ arid passed into linear alpha olefin synthesis
reactor
I~. 5ome of the hydrocarbon synthesis reaction water maybe removed fi'Qrn 'phe
unreacted synthesis gas, to adjust the water vapor pressure u~ 'the linear
alpha
olefin synthesis gas feed stream 58, tv the desired level. Most of the Ca+
hydrocarbons are also removed before the gas is passed into reactor I2.
Reactor
12 and units 14 and 16 are the same ~ for Figure 1, and 'their operation and
.. .
function need not be repeated here. Thus, the linear alpha olef ns produced in
reactor 12 are removed from 16 via line 30_ However, the other CZO+ and Cxp-
hydrocarbons synthesised m 12, and removed from.14 and I6 via lines 24 and .
~~~MEND,E~~S~~ii~E~~ ~~'
cm~,~ ~o;+~nq~la;"~nnd X1:16 ; ~ : ~ 3 «. ~.. yi'~,~ :n~w;3~ P.006
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-23-
32 (as in Figure 1), are passed into line 56, in which they mix with the
liquid
hydrocarbons removed from the hydrocarbon synthesis reactor 52. This mixture
is passed to upgrading operations.
[0034] Turning now to Figure 3, an integrated gas conversion plant 60
comprises a synthesis gas generating unit 62, a hydrocarbon synthesis unit 64,
a
hydrocarbon upgrading unit 66 and a hydrogen separating unit 68, along with a
linear alpha olefin hydrocarbon synthesis reactor 12 and associated linear
alpha
olefin product separation units 14 and 16. Units 12, 14, 16 and associated
flow
'lines are the same as in Figures 1 and 2. Natural gas, oxygen and steam are
fed
into the synthesis gas generating unit 62 via lines 70, 72 and 74,
respectively, to
generate a synthesis gas, from natural gas, comprising a mixture of HZ and CO
having a 2.1:1 mole ratio, by a combination of partial oxidation and steam
reforming. Part o~ this gas is fell into unit 64 via line 76. The hydrocarbon
synthesis unit 64 contains one or more hydrocarbon synthesis reactors
producing
hydrocarbons boiling primarily in the lubricant oil and fuels boiling ranges.
Each reactor contains a hydrocarbon synthesis catalyst such as, for example, a
rhenium-promoted cobalt catalytic component supported on titania. The unit
also contains heat exchangers and separation drums for cooling and separating
out the synthesis reaction water from the C4.,. hydrocarbons in the gas. The
synthesized C4.,. hydrocarbons are passed, via line 78, from unit 64 to hydro-
carbon upgrading unit 66. Unit 66 contains one or more fractionators, as well
as
one or more hydrotreating reactors, such as isomerization units to lower the
pour
point of the paraffinic hydrocarbons. The upgraded hydrocarbons are removed
from 66 via line 67. A synthesis gas slip-stream i$ withdrawn from line 76,
via
line 80, and passed into unit 68. Unit 68 separates hydrogen from the
synthesis
gas to produce a stream of CO-enriched synthesis gas, which is removed from
68, via line 82. The Ha:CO mole ratio in this gas stream is such that, when
mixed with recycled, unreacted synthesis gas from reactor 12, the HZ:CO mole
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-24-
ratio of the feed gas entering 12 is lower than 2:1, preferably lower than
1.8:1
and more preferably about I :1. This combined feed stream is passed into the
alpha olefin synthesis reactor 12 via line 82. Reactor 12 operates at reaction
conditions, as in the embodiment of Figure 1 and includes the low H2:C0 mole
ratio in the feed gas, to satisfy the requirement of the Condition Factor
being
greater than or equal to 50. Reactor 12 produces linear alpha olefins at a
high
yield which, along with the unreacted synthesis gas, is passed to separation
unit
14, via line 20. If the desired alpha olefin fraction falls in the carbon
number
range of 4 to 20, it is p'~eferred to make hydrocarbons in reactor 12 that
boil
primarily in the naphtha and lower fuels boiling ranges. The hydrogen
separated
in unit 68 is passed to the hydrocarbon upgrading unit 66, via line 84.
Unreacted
synthesis gas separated in 14 is passed via line 26, back into line 82 in
which it
mixes with the CO-enriched gas formed in 68. As in Figures I and 2, a portion
of the recycled gas is purged via line 28 and is used as low BTU fuel gas or
reused for synthesis gas generation in 62. The C2i+ hydrocarbons produced in
12, and separated from the C2o_ linear alpha olefins in 14, are removed from
14
via line 24 and passed to line 78, for further processing in unit 66. The
linear
alpha olefin product of reactor 12 is removed from unit 16 via line 30. Line
32
passes the CZO_ hydrocarbons separated from the C2o_ linear alpha olefins to
further processing for naphtha and low BTU value fuels.
[0035] At least a portion of the hydrocarbons produced by a hydrocarbon
synthesis process, including the saturates produced by the linear alpha olefin
production process according to the invention, are upgraded to more valuable
products. The upgrading comprises one or more of fractionation and conversion.
By conversion is meant one or more operations in which the molecular structure
of at least a portion of the hydrocarbon is changed and typically in the
presence
of a catalyst. Hydrotreating is a conversion in which hydrocarbons are reacted
with hydrogen in the presence of a catalyst and includes, for example, hydro-
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isomerization, hydrocracking, hydrodewaxing, hydrorefining, all conducted at
conditions well known in the literature. In the context of the invention,
conversion also refers to chemically treating the linear alpha olefins to
separate
them from the other hydrocarbons in the separated C4-CZO fraction and treating
the separated -material to convert it back to the desired linear alpha olefin
product, by any suitable means, including the etherification and metal
alkylation
procedures mentioned above.
[0036] The invention will be further understood with reference to the
examples below.
EXAMPLES
[0037] In all of the examples, the reactors were downflow, isothermal tube
i
reactors containing a fixed bed of particulate catalyst (particle size below
80
mesh), held in place by quartz wool. The catalyst bed was diluted with below
80
mesh quartz or SiC to ensure isothermal conditions. In all cases they were
operated at a total pressure of 20 atmospheres (2000 kPa), with a synthesis
gas
flow rate of either 6,000 or 40,000 V/V/hr. The synthesis gas fed into the
reactor
comprised of HZ and CO as the reactants, along with Ar and/or Ne as a diluent
and internal standard for establishing mass balances. An He diluent was also
used to balance the total pressure of the reactants and internal standards.
The
mass balances were calculated from the feed flow rates combined with feed and
product compositions obtained by gas chromatography (GC) or GC-mass-
spectrometric (GC-MS) analysis. CO conversions were calculated by known
methods of using internal standards. The Hewlett'-Packard 5890 GC and Balzers
TGG 300 MS coupled to the Hewlett-Packard 5890 GC (GC-MS mode) were
calibrated using certified gas blends. The Ha to CO mole ratio varied from 1:1
to
more than 2:1. The CO conversion was adjusted in all experiments to the target
12% level by adjusting the feed rate. In two of the runs, those of Examples 7
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-26-
and 8, the catalyst comprised rhenium-promoted cobalt on titanic (11 ~wt% Co,
1
wt% Re), of the type used for Fischer-Tropsch hydrocarbon synthesis in a
slurry
synthesis reactor. The rest of the examples, Examples 1-6, used an
unsupported,
rhenium-promoted cobalt catalyst. This catalyst was prepared by impregnating
cobalt powder with an activating solution of cobalt nitrate and perrhenic
acid.
The impregnation solution was prepared by dissolving 26.3 g of Co(N03)2 and
2.3 g perrhenic acid (54 wt% Re) in 9.4 g of distilled and deionized water. An
amount of 34.8 g of this solution was slowly added, with constant mixing, to
50.2 g of 2 micron size''cobalt powder. The exothermic reaction was controlled
by the slow addition rate and stirring. The addition took place over thirty
minutes. The treated cobalt powder was then dried in air at 60°C for
four hours.
The dry powder was then calcined using 1 vol% oxygen in dry nitrogen, by
slowly romping, at 2°C/min to 300°C and then held at
300°C for one hour. After
cooling, the dry powder was loaded into the reactor for the hydrocarbon
synthesis runs. The conditions and results of the eight runs are summarized in
the table below.
Example 1
[0038] In this experiment, the reactor was operated at 210°C, the mole
ratio
of the HZ to CO in the synthesis gas feed was 1:1 and the unsupported cobalt
catalyst described above was used for the synthesis. The partial pressure of
both
the H2 and CO in the reactor was 500 kPa. The CF value in this experiment was
59.
Example 2
[0039] In this experiment, the effect of temperature was tested by operating
the reactor at 221°C. All other conditions were the same as in Example
1. The
CF value in the experiment was 52.
CA 02507722 2005-05-30
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Example 3
[0040] In this experiment, the effect of the H2:C0 ratio was tested. The
reactor was operated at the same conditions as in Example l, except for the
feed
HZ and CO partial pressures, which were 500 kPa and 250 kPa, respectively.
The CF value for the run was 51.
Example 4
[0041] In this experiment, the effect of the total synthesis gas pressure was
tested. The reactor was operated at the same conditions as in Example 1,
except
for the partial pressures of H2 and CO in the reactor feed , which were both
only
250 kPa instead of the 500 kPa in Example 1. The CF value was 59.
i
Example 5
[0042] The effect of H20 vapor in the feed gas was demonstrated in this
experiment. The reactor was operated as in Example 3, except that in this run
400 kPa steam was added to the feed gas by replacing an equal concentration
(partial pressure) of the inert diluent in the feed. The CF value in the
experiment
was 63.
Comparative examples 6, 7, and 8 (outside the scope of the invention)
Example 6
[0043] This experiment demonstrates that reduced CO conversion alone. will
not ensure high olefin selectivity, if the CF value is not above 50. The
respective partial pressures of the HZ and CO in the reactor were 1000 kPa and
500 kPa, providing a HZ:CO ratio of 2:1 in the dry feed gas. The CO conversion
CA 02507722 2005-05-30
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_7$_
was set to 12%, just as in Examples 1-5. The temperature was 220°C. The
overall CF value in this experiment was 45.
Example 7
[0044] In this experiment, the catalyst was different from those used in
Examples 1-6 and comprised the rhenium-promoted cobalt (11 wt% Co, 1 wt%
Re) catalytic component supported on titania. The reactor was operated at the
same conditions as in Example 6: 220°C, the mole ratio of the H2 to CO
in the
synthesis gas feed was 2:1 and the respective partial pressures of the HZ and
CO
in the reactor were 1000 kPa and 500 kPa, respectively. Consequently, the CF
value in the experiment was also the same as in Example 6, i.e., 45.
Example 8
[0045] The same rhenium-promoted cobalt on titania catalyst used in
Example 7 was also used in this experiment. The reactor was operated at a
similar temperature of 219°C, the mole ratio of the H2 to CO in the
synthesis gas
feed was about 2.2:1 and the respective partial pressures of the H2 and CO in
the
reactor were 1420 kPa and 650 kPa. No H20 vapor was used in the feed gas.
However, while the CO conversion for Examples 1-7 was low, only 12%, in this
experiment it was closer to typical hydrocarbon synthesis conditions, namely
61 %. The CF value in this experiment was the lowest, 15.
CA 02507722 2005-05-30
WO 2004/060836 PCT/US2003/036896
ExpCatalyst Temp.CO PH2 Pco CF* PHZO1-alkene 1-octene/
conv. in C4-Clooctane
C % kPa kPa kPa %
1 Co 210 12 500 500 59 0 61 1.92
2 Co , 221 12 500 500 52 0 62 1.76
3 Co 210 12 500 250 51 0 55 1.18
4 Co 210 12 250 250 59 0 64 1.94
Co 210 12 500 250 63 400 66 1.99
6 Co 220 12 1000 500 45 0 26 0.32
7 Co-Re/Ti02220 12 1000 500 45 0 35 0.34
f Co-Re/Ti022-T9 61 1420 650 ~15 0 18 0.13
*Note: CF = 200 - 0.6T + 0.03PHZO - 0.6Xco - 8(H2:C0)
where,
T = average reactof temperature ~n °C
PH20 = p~'hal pressure of the water in the synthesis gas feed to the reactor
in kPa
Xco = CO conversion, expressed as percent
H2:C0 = Hydrogen to CO molar ratio in the synthesis gas feed to the reactor
[0046] In Experiments 1-5, all of which fall within the scope of the
invention,
oxygenate selectivity was negligible, in that no oxygenates were detected in
the
synthesized C4-C1o hydrocarbon products by gas chromatography using a flame-
ionization detector. C02 selectivities were also negligible in all runs,
typically
around 0.5 % or less, clearly affirming the non-shifting nature of the cobalt
catalysts used.
[0047] Referring now to the data in the Table, it is seen that Examples 1 and
2 were both run at an HZ:CO mole ratio of 1:1, with the partial pressure of
both
the H2 and CO being 500 kPa, but at different temperatures. Comparing the
results for these two runs reveals that both temperatures produced about the
same amount of linear alpha olefins in the C4-Clo hydrocarbon fraction
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synthesized in the reactor, suggesting a relatively small effect of
temperature on
the alpha olefin selectivity. The lower synthesis temperature of 210°C
produced
a slightly greater selectivity for C8 linear alpha olefins. Overall, however,
this
comparison shows that the reaction temperature can be used for adjusting the
Fischer-Tropsch synthesis alpha and thus increasing'the total alpha olefin
yield,
particularly in the lower boiling range.
[0048] The Example 3 results reveal that increasing the H2:C0 mole ratio
from 1:1 up to 2:1 resulted in a significant decrease in the linear alpha
olefin
content of the synthesized C4-Clo hydrocarbon fraction, from about 62% down to
55%, with an even greater decrease in C$ olefin selectivity from 1.92 down to
1.18. The reduced alpha olefin yield and selectivity correlates well with the
borderlir[e CF value of 51. This example also clarifies that low CO conversion
~ '
alone is an insufficient condition for high alpha olefin selectivity and
yield. On
the other hand, the value Condition Factor a of the invention can define the
operating conditions and predict the high alpha olefin yield.
[0049] A comparison of the Example 4 results with those. of Example 1,
which differ only by the pressure of the H2 and CO reactants, shows that
changing the partial pressures of both the H2 and CO reactants in the feed
does
not adversely effect either the total fraction of linear alpha olefins in the
C4-Clo
hydrocarbon fraction or the selectivity for C8 linear alpha olefins. Comparing
the results from Examples 5 and 3 clearly demonstrates the beneficial effect
of
adding Hz0 vapor to the feed gas. The 400 kPa water in the feed gas
compensated for the higher H2:C0 mole ratio of 2:1. Thus, the linear alpha
olefin content of the C4-Clo hydrocarbon fraction jumped from the 55% results
of Example 3, up to 66% in the run of Example 4. At the same time, the
selectivity for the C8 linear alpha olefin increased from 1.18 up to 1.99, an
almost two-fold increase. Again, this example demonstrates that the selection
of
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the preferred conditions for making linear alpha olefins is not obvious and
that
selecting operating conditions which result in a Condition Factor of the
invention having a value greater than or equal to 50, can provide high yields
of
the desired linear alpha olefins.
[0050] Comparative Examples 6-8 were all run at the higher reaction
temperature of about 220°C and the less desirable, higher 1-i2:C0
reactant mole
ratio of at least 2:1, at both low (12%) and high (61%) CO conversion. These
runs demonstrate that the alpha olefin selectivity abruptly drops when the
Condition Factor falls below the critical value of 50. They also compare the
unsupported rhenium-promoted Co catalyst with the rhenium-promoted cobalt
(11 wt% Co and 1 wt% Re) on titania catalyst used for Fischer-Tropsch
hydrocarbon production. At the same CO conversion level of 12%, the
supported Fischer-'2'ropsch hydrbcarbon synthesis catalyst of Example 7
produced more linear alpha olefins than did the unsupported Re-Co catalyst of
Example 6. This has two has consequences. The first is that a conventional,
supported Fischer-Tropsch catalyst can be used for linear alpha olefin
production. The second is that the reaction and feed conditions in a Fischer-
Tropsch reactor, that is on-line synthesizing distillate fuel and lubricant
oil
hydrocarbons with a supported cobalt catalyst, can be adjusted to increase
linear
alpha olefin production without taking the reactor off line to change the
catalyst,
and vice-versa. Thus, if a Fischer-Tropsch hydrocarbon synthesis process plant
is in existence or planned, another synthesis unit for producing linear alpha
olefins can be included, without the need for a separate unit or plant for
producing the hydrogen and carbon monoxide gas mixture. Linear alpha olefin
synthesis can be accomplished by setting the operating conditions so that the
earlier value of the Condition Factor is greater than or equal to 50.
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[0051] The worst alpha olefin selectivity was obtained in Example 8, in
which the total partial pressure of the H2 and CO reactants was 2070 kPa, the
HZ:CO mole ratio was about 2.2:1 and the CO reactant conversion rate was a
much higher 61 %. Example 8 represents conventional hydrocarbon synthesis
conditions aimed at producing fuels and lubricating oil base stocks, and the
CF
value of 15 is outside the scope of the current invention. Comparing the
Example 7 and Example 8 results' shows that the olefin selectivity at the
lower
CO conversion level of 12% was about twice that obtained at the higher CO
conversion level of 61% (35% vs. 18%). However, Example 7 also
demonstrates that while lowering CO conversion is preferred, lower CO
conversion alone is not sufficient to achieve an effective level of alpha
olefin
production. This fact is reflected in the lower than acceptable CF in the runs
of,
Examples 7 and 8.
